Process for the utilization of refinery C4 streams

ABSTRACT

A process is disclosed for preparing a C 4  stream for feeding to an alkylation process which reacts isobutane with butene to produce isooctane. The C 4  stream is treated in a first distillation column reactor to remove dienes and mercaptans and separate out any C 5 &#39;s which might be present. The treated C 4 &#39;s are then fed to a second distillation column reactor that concurrently isomerizes 1-butene to 2-butene and splits the normal C 4 &#39;s from the iso C 4 &#39;S. The iso C 4 &#39;s are then fed to a third distillation column reactorwhere a portion of the isobutene is saturated to isobutane. The C 4 &#39;s from the isomerization/splitter are combined with the C 4 &#39;s from the hydrogenation unit and fed to a cold acid alkylation unit. The third distillation column may also oligomerize a portion of the isobutene to diisobutene in the upper end which is saturated in the bottom of the column to isooctane.

BACKGROUND OF THE INVENTION

[0001] 1. Field of the Invention

[0002] The present invention relates to a process for the utilization ofrefinery C₄ streams containing iso and normal butanes and butenes. Moreparticularly the invention relates to a process for producing alkylatefeed from a mixed C₄ stream. In one embodiment the invention relates toa process which produces isooctane in addition to alkylate feed.

[0003] 2. Related Information

[0004] Refinery C₄ streams have recently been utilized in the productionof methyl tertiary butyl ether (MTBE) for use as an oxygenate additiveand octane improver in motor gasolines. Processes and catalyst systemsfor their use in this manner have been developed over the years andinclude U.S. Pat. Nos. 4,215,011; 4,242,530; 4,232,177; 4,307,254;4,336,407; 4,375,576 and like patents.

[0005] Refinery C₄ streams have also been used as the source of butenesand isobutane for feed to cold acid alkylation processes which produceisooctane. Generally there has been an imbalance in the amount ofisobutene and butenes in refinery C₄ streams which have led to the useof the MTBE processes which utilize the isobutene, effectively removingit from the C₄ stream.

[0006] Environmental concerns have led at least the state of Californiato ban the use of MTBE in gasoline. Other processes for the balancing ofthe iso/normal butenes are now required.

SUMMARY OF THE INVENTION

[0007] Briefly, the present invention is an integrated process for thepreparation of paraffin alkylate in which a C₄ hydrocarbon feed is firsttreated to remove dienes and mercaptans, for example, by reacting thedienes and mercaptans to form sulfides, separating the C₄'s from heavymaterial comprising the sulfides, for example, by fractionation. Thetreated C₄ feed is then subjected to isomerization to convert butene-1to butene-2 and the iso C₄'s are separated from the normal C₄, forexample by fractionation. The iso C₄ portion is hydrogenated to convertisobutene to isobutane and the C₄ fractions reunited and subjected toparaffin alkylation to produce an alkylate comprising isooctane.

[0008] In a preferred embodiment the C₄ stream is first treated in afirst distillation column reactor to remove dienes and mercaptans andseparate out any C₅'s which might be present. The treated C₄'s are thenfed to a second distillation column reactor that concurrently isomerizes1-butene to 2-butene and splits the normal C₄'S from the iso C₄'S. Theiso C₄'s are then fed to a third distillation column reactor where aportion of the isobutene is saturated to isobutane. The C₄'s from theisomerization/splitter are combined with the C₄'S from the hydrogenationunit and fed to a cold acid alkylation unit.

[0009] In one embodiment the third distillation column reactor alsooligomerizes a portion of the isobutene to diisobutene in the upper endwhich is saturated in the bottom of the column to isooctane. A portionof the unreacted isobutene is also hydrogenated to isobutane. Theremainder of the iso C₄'s are fed to alkylation unit.

BRIEF DESCRIPTION OF THE DRAWING

[0010]FIG. 1 is a flow diagram in schematic form of one embodiment ofthe invention.

[0011]FIG. 2 is a flow diagram in schematic form of a second embodimentof the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENT

[0012] Referring now to FIG. 1 a flow diagram of one embodiment of theinvention is shown. Feed to a cold acid alkylation unit must be low indienes and mercaptans. To this end the mixed C₄ feed in flow line 101 istreated in a distillation column reactor 10 containing a bed 12thioetherification catalyst which reacts the mercaptans with dienes inthe feed to produce higher boiling sulfides which are removed in thebottoms in flow line 103 along with any C₅'S in the feed. Alternativelythis column may be operated to send the C₅'s overhead with the C₄'s andto recover the C₅'S in the bottoms from column 20 along with the nC₄'sin line 105. Catalysts which are useful for this reaction include theGroup VIII metals, such as palladium and nickel. Generally the metalsare deposited as oxides on an alumina support. The supports are usuallysmall diameter extrudates or spheres. A suitable catalyst for thereaction is 58 wt. % Ni on 8 to 14 mesh alumina spheres, supplied byCalcicat, designated as E-475-SR. Typical physical and chemicalproperties of the catalyst as provided by the manufacturer are asfollows: TABLE I Designation E-475-SR Form Spheres Nominal size 8 × 14Mesh Ni wt. % 54 Support Alumina

[0013] The hydrogen rate to the reactor, which is fed along with themixed C₄ feed in flow line 101, must be sufficient to maintain thereaction which is understood to be the “effectuating amount of hydrogen”as that term is used herein, but kept below that which would causeflooding of the. Generally, the mole ratio of hydrogen to diolefins andacetylenes in the feed is at least 1.0 to 1.0, preferably at least 2.0to 1.0 and more preferably at least 10 to 1.0.

[0014] The catalyst also catalyzes the selective hydrogenation of thepolyolefins contained within the feed and to a lesser degree theisomerization of some of the mono-olefins. Generally the relativeabsorption preference is as follows:

[0015] (1) sulfur compounds

[0016] (2) diolefins

[0017] (3) mono-olefins

[0018] If the catalyst sites are occupied by a more strongly absorbedspecies, reaction of these weaker absorbed species cannot occur.

[0019] The reaction of interest is the reaction of the mercaptans and/orhydrogen sulfide (H₂S) with diolefins. In the presence of the catalystthe mercaptans will also react with mono-olefins. However, there is anexcess of diolefins to mercaptans and/or hydrogen sulfide (H₂S) in thefeed and the mercaptans preferentially react with them before reactingwith the mono-olefins. The equation of interest which describes thereaction is:

[0020] Where R, R₁ and R₂ are independently selected from hydrogen andhydrocarbyl groups of 1 to 20 carbon atoms. If there is concurrenthydrogenation of the dienes, then hydrogen will be consumed in thatreaction. The only mercaptans expected to be present in the mixed C₄feed are the lower boiling ones such as methyl mercaptan.

[0021] The treated feed is taken overheads via flow line 102 and thenfed to a second distillation column reactor 20 which contains a bed 22of isomerization catalysts. In the second distillation column reactor 20the isobutene, isobutane and unreacted 1-butene are concurrentlyseparated from normal butanes and both isomers of 2-butene. Concurrently1-butene is isomerized to 2-butene to help in the separation.Essentially cis and trans 2-butene and normal butane are removed asbottoms via flow line 105 while the remaining C₄'s (isobutane andisobutene) are removed as overheads via flow line 104. The 2-butenes arepreferred over 1-butenes for alkylation, since the 2-butene alkylateproduct has a higher octane number.

[0022] The distillation column reactor is generally operated at overheadtemperatures in the range of 80 to 180° F., more preferably 100 to 150°F. at pressures in the range of 50 to 110 psig (bearing in mind theeffect of pressure on temperature as discussed above). In its morepreferred embodiments the present process is operated under conditions,particularly temperature and pressure, which tend to exclude butene-2from contact with the catalyst while holding the butene-1 in contactwith the catalyst. Thus, as butene-1 is isomerized to butene-2 it dropsdown in the column away from the catalyst and is removed as bottoms.

[0023] A reflux is preferably included in the system. The reflux ratiocould vary over the rate 0.5:1 to 33:1. In practice, the higher ratiomay be used to compensate for a short catalyst bed such as required forexperimental work. In commercial size units the catalyst bed would beprovided so that lower reflux and hence higher unit productivity couldbe obtained at lower operating cost. In a modification designed tooptimize the operation of the embodiments of both FIG. 1 and FIG. 2 thecatalyst bed 22 may be divided into several smaller zones withconventional distillation staging between each zone (not shown).

[0024] A catalyst suitable for the present process is 0.5% PdO on ⅛″Al₂O₃ (alumina) extrudates, hydroisomerization catalyst, supplied byEngelhard Industries. The catalyst is believed to be the hydride ofpalladium which is produced during operation. The hydrogen rate to thedistillation column reactor must be sufficient to maintain the catalystin the active form because hydrogen is lost from the catalyst byhydrogenation. The hydrogen rate must be adjusted such that there issufficient hydrogen to replace hydrogen lost from the catalyst which isunderstood to be the “effectuating amount of hydrogen” as that term isused herein but kept below that required for hydrogenation of butenes orto cause flooding of the column. Generally, the mole ratio of hydrogento C₄ hydrocarbon fed to the bed of the present invention will be about0.01 to 0.60, preferably 0.01 to 0.10.

[0025] Another suitable catalyst for the reaction is 0.34 wt % Pd on 7to 14 mesh Al₂O₃ (alumina) spheres, supplied by United Catalysts Inc.designated as G-68C. Typical physical and chemical properties of thecatalyst as provided by the manufacturer are as follows: TABLE IIDesignation G-68C Form Sphere Nominal size 7 × 14 mesh Pd. wt % 0.3(0.27-0.33) Support High purity alumina

[0026] Finally the separated isobutene (along with trace amounts of1-butene and isobutane) are fed via line 104 to a third distillationcolumn reactor 30 above a bed 32 of hydrogenation catalyst where aportion of the isobutene is saturated to isobutane. Hydrogen for thereaction is fed below the bed via flow line 106. The Pd and Ni catalystsare useful for this hydrogenation also. A distillation is carried out tohave a condensing liquid in the column which the hydrogen is occludedwhich improves the contact between the hydrogen, catalyst andhydrocarbon. Thus an overhead is taken via flow line 107 with thecondensible material being condensed in condenser 34 and returned to thedistillation column reactor 30 as reflux via flow line 108. Vapors areremoved via flow line 110

[0027] The bottoms from this distillation column reactor in flow line109 along with the bottoms from the separation/isomerization column inflow line 105 are fed to a cold acid alkylation process. Referring nowto FIG. 2 a second embodiment of the invention is shown with the samereference numerals depicting the same items as in FIG. 1. The first twodistillation column reactors 10 and 20 are identical to that of FIG. 1.The third distillation column reactor 30, however has a bed of acidiccation exchange resin in an upper bed 33 which oligomerizes a portion ofthe isobutenes to diisobutene which is then hydrogenated in a lower bedof hydrogenation catalyst 32 along with a portion of the isobutene. Thehydrogenated diisobutene is 2,2,4-trimethyl pentane or isooctane of 100octane number and is taken as bottoms via flow line 109. The overheadsproduct containing the iC₄'s and iC₄='s are taken via flow line 111 andfed to the cold acid alkylation unit (not shown) along with the bottomsfrom the second distillation column reactor 20.

[0028] Preferably the hydrogenation in column 30 is carried out asdescribed in copending U.S. patent application Ser. No. 09/474,192 filedDec. 29, 1999, which is incorporated herein in its entirety, byconcurrently passing the feed containing diisobutene and hydrogendownflow through a reaction zone containing a hydrogenation catalyst ata pressure of less than 300 psig pressure, preferably less than 275psig, for example less than 200 psig, and for example at least about 100psig at a temperature within the range of 300° F. to 700° F. to producean effluent, said temperature and pressure being adjusted such that thetemperature of the effluent is above its boiling point and below its dewpoint, whereby at least a portion but less than all of the material insaid reaction zone is in the vapor phase. Preferably the weight hourlyspace velocity (WHSV), i.e., the weight of petroleum feed per hour pervolume of catalyst is greater than 6 hr⁻¹, preferably greater than 8hr⁻¹ and more preferably greater than 15 hr¹.

[0029] The preferred alkylation process comprises alkylation ofisoparaffin with olefin comprising contacting a fluid system comprisingacid catalyst, isoparaffin and olefin in concurrent flow, preferablydownflow into contact in a reaction zone with internal packing, such as,a coalescer under conditions of temperature and pressure to react saidisoparaffin and said olefin to produce an alkylate product. Preferably,the fluid system comprises a liquid and is maintained at about itsboiling point in the reaction zone.

[0030] The reaction may be carried out in an apparatus comprising avertical reactor containing a coalescer in the reaction zone, which maycomprise the entire column or a portion thereof.

[0031] The process is more completely described in co-owned patentapplication having docket number CDT 1769/79 (U.S. S No. 60/323,227filed Sep. 14, 1901) which is hereby incorporated by reference.

[0032] The preferred alkylation process employs a downflow reactorpacked with contacting internals or packing material (which may be inertor catalytic) through which passes a concurrent multi phase mixture ofsulfuric acid, hydrocarbon solvent and reactants at the boiling point ofthe system. Adjusting the pressure and hydrocarbon composition controlsthe boiling point temperature. The reactor is preferentially operatedvapor continuous but may also be operated liquid continuous. Thepressure is preferentially higher at the top of the reactor than at thebottom. Adjusting the flow rates and the degree of vaporization controlsthe pressure drop across the reactor. Multiple injection of olefin ispreferred. The product mixture before fractionation is the preferredcirculating solvent. The acid emulsion separates rapidly from thehydrocarbon liquid and is normally recycled with only a few minutesresidence time in the bottom phase separator. Because the products arein essence extracted from the acid emulsion, the reaction and/oremulsion promoters may be added without the usual concern for breakingthe emulsion. The process may be described as being hydrocarboncontinuous as opposed to acid continuous.

[0033] The coalescer comprises a conventional liquid-liquid coalescer ofa type which is operative for coalescing vaporized liquids. These arecommonly known as “mist eliminators” or “demisters”. A suitablecoalescer comprises a mesh such as a co-knit wire and fiberglass mesh.For example, it has been found that a 90 needle tubular co-knit mesh ofwire and fiberglass such as manufactured by ACS Industries LLC ofHouston, Tex., can be effectively utilized, however, it will beunderstood that various other materials such as co-knit wire and Teflon(Dupont™), steel wool, polypropylene, PVDF, polyester or various otherco-knit materials can also be effectively utilized in the apparatus.

The invention claimed is:
 1. An integrated process for the preparationof paraffin alkylate from a C₄ feed containing isobutane, isobutene,normal butane, butene-1, butene 2, dienes and mercaptans comprising:removing dienes and mercaptans from said C₄ feed; isomerizing a portionof the butene-1 to butene-2; separating iso C₄ components from thenormal C₄ components; hydrogenating a portion of the isobutene in saidiso C₄ components to isobutane; recombining said normal C₄ componentsand said iso C₄ components and alkylating isobutane and normal butenesin said recombined C₄ components to produce an alkylate comprisingisooctane.
 2. The process according to claim 1 wherein said removing ofdienes and mercaptans comprises reacting said dienes and mercaptans inthe presence of a thioetherification catalyst and hydrogen underthioetherification conditions to form sulfides and fractionating theresultant mixture to separate a heavy portion comprising said sulfides.3. The process according to claim 1 wherein said isomerizing is carriedout in the presence of an isomerization catalyst and hydrogen underisomerization conditions.
 4. The process according to claim 1 whereinsaid separating is by fractionation.
 5. The process according to claim 1wherein said alkylating is carried out in the presence of an acidcatalyst under alkylation conditions.
 6. An integrated process for thepreparation of paraffin alkylate from a C₄ feed containing isobutane,isobutene, normal butane, butene-1, butene 2, dienes and mercaptanscomprising: reacting said dienes and mercaptans in the presence of athioetherification catalyst and hydrogen under thioetherificationconditions to form sulfides and fractionating the resultant mixture toseparate a heavy portion comprising said sulfides; isomerizing a portionof the butene-1 to butene-2 in the presence of an isomerization catalystand hydrogen under isomerization conditions; separating iso C₄components from the normal C₄ components by fractionation; hydrogenatinga portion of the isobutene in said iso C₄ components to isobutane;recombining said normal C₄ components and said iso C₄ components andalkylating isobutane and normal butenes in said recombined C₄ componentsin the presence of an acid catalyst under alkylation conditions toproduce an alkylate comprising isooctane.
 7. A process for theutilization of refinery C₄ streams in the production of gasolinecomprising the steps of: (a) feeding hydrogen and a mixed C₄ streamcontaining normal butane, isobutane, 1-butene, 2-butene, isobutene,dienes, mercaptans and C₅'s to a first distillation column reactorcontaining a bed of thioetherification/hydrogenation catalyst; (b)concurrently in said first distillation column reactor, (i) reacting themercaptans and a portion of the dienes in the presence of saidthioetherification/hydrogenation catalyst to produce sulfides, (ii)reacting at least a portion of said dienes with said hydrogen to formmono olefins including additional butenes, and (iii) separating saidsulfides and said C₅'s from said normal butane, isobutane, 1-butene,2-butene and isobutene by fractional distillation; (c) removing saidC₅'s and said sulfides from said first distillation column reactor as afirst bottoms; (d) removing said normal butane, isobutane, 1-butene,2-butene and isobutene from said first distillation column reactor as afirst overheads; (e) feeding said first overheads containing said normalbutane, isobutane, 1-butene, 2-butene and isobutene to a seconddistillation column reactor containing an bed of isomerization catalyst;(f) concurrently in said second distillation column reactor, (i)isomerizing a portion of the 1-butene to 2-butene, and (ii) separatingthe 2-butene and the normal butane from the isobutane, isobutene andunreacted 1-butene; (g) removing the 2-butene from said seconddistillation column reactor as a second bottoms; (h) removing theisobutane, unreacted 1-butene and isobutene from said seconddistillation column reactor as a second overheads; (i) feeding saidhydrogen and second overheads containing said normal butane, isobutane,isobutene and 1-butene to a third distillation column reactor containinga bed of hydrogenation catalyst to concurrently; (i) hydrogenate aportion of the 1-butene and isobutene to form a reaction productcomprising butane and isobutane and (ii) fractionate the reactionproduct to produce a third overheads; and (j) removing the normalbutane, isobutane, 1-butene an isobutene from said third distillationcolumn reactor as a third bottoms.
 8. The process according to claim 7wherein said second and third bottoms are fed to a cold acid alkylationunit.
 9. The process according to claim 7 comprising recovering saidthird overheads comprising hydrogen, isobutane, 1-butene and isobutene;condensing said third overheads to recover a condensate comprisingisobutane, 1-butene and returning said condensate to said thirddistillation column reactor as reflux.
 10. The process according toclaim 7 wherein said third distillation column reactor contains a bed ofacidic cation exchange resin above said bed of hydrogenation catalystand a portion of the isobutene is oligomerized to produce diisobutenewhich is hydrogenated in said bed of hydrogenation catalyst and saiddiisobutene is removed as a third bottoms and said isobutane, 1-buteneand isobutene are removed as a third overheads.
 11. The processaccording to claim 10 wherein said second bottoms and said thirdoverheads are fed to a cold acid alkylation unit.
 12. The processaccording to claim 10 wherein the normal butane, isobutane, 1-butene andisobutene is condensed and a portion is returned to said thirddistillation column reactor as reflux.
 13. The process according toclaim 7 wherein said hydrogenation is downflow.